Thermally integrated process for the production of ethylene oxide from a flow of ethanol

ABSTRACT

Process for dehydrating an ethanol feed to ethylene then oxidating the ethanol to ethylene oxide, comprising vaporizing the feed and a flow of diluting water comprising recycled ethanol to produce a vaporized feed, dehydrating a mixture of vaporized feed and a vaporized flow of diluting water comprising ethanol, separating the dehydration effluent into ethylene and water effluents, purifying at least a portion of the water effluent and separating a flow of treated water and a flow of diluting water comprising ethanol, recycling and vaporizing at least a portion of the separated diluting water by partial or complete vaporization in an exchanger using heat exchange with a quench flow obtained from the oxidation, said quench flow, cooled, then recycled to the oxidation reactor(s), and oxidation of the ethylene in the effluent comprising ethylene into ethylene oxide, this oxidation using a tubular oxidation reactor cooled by vaporization of said quench flow.

FIELD OF THE INVENTION

The present invention relates to a process for the transformation ofethanol to ethylene oxide including a thermally integrated step fordehydration of ethanol and a thermally integrated step for oxidation ofthe ethylene produced.

PRIOR ART

The ethanol to ethylene dehydration reaction has been known in detailsince the end of the 19^(th) century. “The Dehydration of Alcohols overAlumina. I:The reaction scheme”, H. Knözinger, R. Köhne, Journal ofCatalysis (1966), 5, 264-270 is considered to be the baselinepublication regarding studies of alcohol dehydration, including that ofethanol. This reaction is known to be highly endothermic, equilibratedand displaced towards ethanol at high temperatures. The drop intemperature corresponding to the total conversion of pure ethanol in anadiabatic reactor is 380° C. At lower temperatures, the ethanol isconverted into diethyl ether (DEE). This reaction “intermediate” may bepresent in ethylene dehydration reactions in which the conversion ispartial, or between two reactors in multi-reactor processes. DEE maythen be converted into ethylene at a higher temperature. The referencecatalyst which is frequently used is a monofunctional acid catalyst;gamma alumina is the most cited catalyst. Zeolites are also used forthis application, in particular ZSM-5 since the 1980s—see, for example,“Reactions of ethanol over ZSM-5”, S. N. Chaudhuri & al., Journal ofMolecular Catalysis 62:289-295 (1990).

U.S. Pat. No. 4,232,179 describes a process for the dehydration ofethanol to ethylene in which the heat necessary for the reaction issupplied by introducing a heat transfer fluid mixed into the reactorwith the feed. The heat transfer fluid is either steam from an externalsource or an external flow originating in the process, or a recycle of aportion of the effluent from the dehydration reactor, i.e. ethylenewhich is produced. Introducing a mixture of the feed with said heattransfer fluid means that the heat necessary to maintain the temperatureof the catalytic bed at a level compatible with the desired conversionscan be provided. In the case in which the heat transfer fluid is theeffluent from the dehydration reactor, a compressor for recycling saideffluent is necessary. However, recycling the ethylene produced by thereaction constitutes a disadvantage, as the introduction of ethylenemodifies the equilibrium of the dehydration reaction. In addition,ethylene participates in secondary oligomerization reactions, hydrogentransfer reactions and olefin disproportionation reactions; these arereactions with an order of more than 0 with respect to their reagent.Increasing the concentration of ethylene right from the start of thereaction multiplies the formation of by-products. The loss of ethyleneis thus higher, which results in a drop in selectivity.

Patent application WO 2007/134415 describes a process for thedehydration of ethanol to ethylene which has been improved compared withU.S. Pat. No. 4,232,179 in order to reduce the investment costs due to areduced amount of equipment and reduced operational costs because steamexternal to the process is not used. In that process, at least a portionof the effluent from the dehydration reactor (mixture of the ethyleneproduced and steam) and superheated steam obtained from the waterproduced by dehydration of the ethanol and condensed in the reactor areused as the heat transfer fluid and enter the dehydration reactor as amixture with the ethanol. Said patent application is silent as to thepressure conditions to be applied between the ethanol feed and theeffluent with the aim of maximizing heat exchange.

U.S. Pat. No. 4,396,789 also describes a process for the dehydration ofethanol to ethylene, in which the ethanol and steam acting as a heattransfer fluid are introduced into the first reactor at a temperature inthe range 400° C. to 520° C. and at a high pressure in the range 20 to40 atm, such that the effluent produced by the dehydration reaction iswithdrawn from the last reactor at a pressure of no less than 18 atm,said reaction product, i.e. ethylene, being capable of undergoing thefinal cryogenic distillation step after cooling without an intermediatecompression step. Said process is also characterized by an exchange ofheat between said dehydration reaction product and the feed introducedinto the first reactor, said reaction product being used to vaporize thefeed entering the first reactor. The unconverted ethanol, at least aportion of the water formed during the reactions of the process and thewater added for the final wash of the gases are recycled to ensurecomplete conversion of the ethanol.

Patent application WO 2011/002699 discloses a process for thedehydration of an ethanol feed to ethylene, comprising vaporization of amixture of ethanol and water and reaction of this mixture in anadiabatic reactor. That application does not address the problem ofmaximizing the recovery of heat with a view to reducing the energyconsumption of the process.

U.S. Pat. No. 3,119,837 describes a process for the oxidation ofethylene in the presence of a catalyst based on silver. It is based onthe positive effect of methane on the selectivity of the reaction forthe conversion of ethylene to ethylene oxide and gives the operatingconditions for oxidation in the presence of purified oxygen. Theprincipal operating conditions that are given are as follows:temperature in the range 200° C. to 300° C. (preferably) and pressure inthe range 15 to 500 psi absolute (preferably).

Patent EP 0 496 470 B1 describes an oxidation catalyst which may be usedin the process described in U.S. Pat. No. 3,119,837. That patentdescribes the principal characteristics of the catalyst as well as theoperating condition ranges (temperatures) in which it can be used. Thecatalyst described is based on silver and the operating conditions arecompatible with those described in U.S. Pat. No. 3,119,837.

Mike Bruscino, “Scientific Design's Ethanol to Monoethylene GlycolTechnology”, Hydrocarbon World, Vol. 5(2), pp. 15-17 (2010) describes aprocess for the production of ethylene glycol from an ethanolbioresource including a process for the dehydration of ethanol toethylene, a process for oxidation of the ethylene produced thereby and aprocess for the conversion of ethylene oxide to glycol. The processdescribed does not consider recycling water to the dehydration section,nor does it consider thermal integration between the oxidation sectionand the dehydration section.

One aim of the invention is to provide a process for the production ofethylene oxide, comprising a step for dehydration of ethanol to ethyleneand a step for catalytic oxidation of the ethylene produced, these twosteps being integrated with each other in order to reduce the energyconsumption and the water consumption, said reductions being effected bya particular arrangement of the thermal integration and the use ofappropriate operating conditions, in particular matching of theevacuation of the heat generated by the oxidation reaction and the levelof dilution necessary at the inlet to the dehydration step.

SUMMARY AND ADVANTAGE OF THE INVENTION

The invention describes a process for the production of ethylene oxidefrom ethanol. This process consists of the integration of two distinctprocesses: catalytic dehydration of the ethanol, and oxidation of theethylene produced.

The process for the dehydration of an ethanol feed to form ethylenepreferably comprises a step for pre-treatment which reduces the quantityof organic or basic nitrogen contained in said feed and converts afraction of the ethanol to DEE, and a step for vaporization of thepre-treated ethanol feed mixed with at least a portion of a flow ofdiluting water comprising ethanol, in an exchanger by means of anexchange of heat with the effluent obtained from the last dehydrationreactor. The catalytic oxidation process employs at least one multi-tubereactor cooled by vaporization of a heat transfer fluid, said fluidpreferably comprising water, in the shell of the reactor.

Said invention has the advantage over prior art processes of reducingthe consumption of water and of reducing the energy consumptionnecessary for conversion of the ethanol to ethylene oxide.

The ethylene oxide produced is a reaction intermediate in the synthesisof many products, for example glycols, polyols, ethoxylates, glycolethers or ethanolamines.

DESCRIPTION OF THE INVENTION

The invention concerns a process for the dehydration of an ethanol feedto ethylene then oxidation of the ethylene to ethylene oxide,comprising:

a) an optional step for preheating said ethanol feed to a temperature inthe range 70° C. to 130° C. by heat exchange with the effluent obtainedfrom step e);

b) an optional step for pre-treatment of the ethanol feed over an acidicsolid operating at a temperature in the range 70° C. to 130° C. in orderto produce a pre-treated ethanol feed;

c) a step for vaporization of a vaporization feed comprising saidpre-treated ethanol feed and at least a portion of the flow of dilutingwater comprising ethanol recycled in accordance with step h) to anexchanger by means of an exchange of heat with the effluent obtainedfrom the last reactor of step e), said vaporization feed beingintroduced into said vaporization step at a pressure in the range 0.1 to1.4 MPa so as to produce a vaporized feed;

d) a step for compressing and superheating said vaporized feed in acompressor so as to produce a feed which is compressed and at thetemperature for supplying to the dehydration section e);

e) a step for dehydration of a mixture comprising said compressed feedand the vaporized flow of diluting water comprising the ethanol obtainedfrom step h), said mixture having a ratio of water to ethanol in therange 1 to 4 by weight, in at least one adiabatic reactor containing atleast one dehydration catalyst and in which the dehydration reactiontakes place, operating at an inlet temperature in the range 350° C. to550° C. and at an inlet pressure in the range 0.3 to 1.8 MPa;

f) a step for separating the effluent obtained from the last adiabaticreactor of step e) into an effluent comprising ethylene at a pressure ofless than 1.6 MPa and an effluent comprising water;

g) a step for purification of at least a portion of the effluentcomprising water obtained from step f) and separation into at least aflow of treated water and a flow of diluting water comprising ethanol;

h) a step for recycling and vaporization of at least a portion of theflow of diluting water comprising ethanol obtained from step g) bypartial or total vaporization in an exchanger using heat exchange with aquench flow obtained from oxidation step j) in a vaporization exchanger,said quench flow, cooled, then being recycled to the reactor or reactorsof oxidation step j), said flow of at least partially vaporized dilutingwater comprising ethanol being recycled upstream of step e), thenon-vaporized portion of the flow of treated water obtained from step g)being recycled to the inlet to vaporization step c);

i) a step for compression and, preferably, for purification of theeffluent comprising ethylene obtained from step f);

j) a step for oxidation of the ethylene comprised in the effluentcomprising ethylene obtained from step i) into ethylene oxide, thisoxidation step comprising at least one tubular oxidation reactor cooledby vaporization of said quench flow preferably comprising water obtainedfrom step h), said quench flow which has thus been reheated beingrecycled to step h).

Feed

In accordance with the invention, the feed treated in the process is anethanol feed.

Said ethanol feed is advantageously a concentrated ethanol feed. Theterm “concentrated ethanol feed” means an ethanol feed comprising apercentage by weight of ethanol of 35% by weight or more. Preferably,said concentrated ethanol feed comprises a percentage by weight ofethanol in the range 35% to 99.9% by weight.

The ethanol feed comprising less than 35% by weight of ethanol may beconcentrated using any means known to the skilled person, for example bydistillation, absorption or pervaporation.

In addition to water, said ethanol feed also advantageously comprises aquantity of alcohols other than ethanol, such as methanol, butanoland/or isopentanol for example, of less than 10% by weight, preferablyless than 5% by weight, a quantity of oxygen-containing compounds otherthan alcohols such as ethers, acids, ketones, aldehydes and/or esters,of less than 1% by weight and a quantity of nitrogen and sulphur,organic and mineral, of less than 0.5% by weight, the percentages byweight being expressed with respect to the total weight of said feed.

The ethanol feed treated in the process of the invention is optionallyobtained by a process for the synthesis of alcohol from fossil fuelssuch as from coal, natural gas or carbonaceous waste, for example.

Said feed may also advantageously derive from non-fossil sources.Preferably, the ethanol feed treated in the process of the invention isan ethanol feed produced from renewable source material obtained frombiomass, often known as “bioethanol”. Bioethanol is a feed produced by abiological pathway, preferably by fermentation of sugars obtained, forexample, from cultures of sugar-containing plants such as sugar cane(saccharose, glucose, fructose and sucrose), beet, or from amylaceousplants (starch) or from lignocellulosic biomass or from hydrolysedcellulose (mainly glucose and xylose, galactose), containing variablequantities of water.

For a more complete description of conventional fermenting processes,reference may be made to the publication “Les Biocarburants, Etat deslieux, perspectives et enjeux du developpement [Biofuels—state of play,perspectives and challenges for development], Daniel Ballerini, EditionsTechnip”.

Said feed may also advantageously be obtained from synthesis gas.

Said feed may also advantageously also be obtained by hydrogenation ofthe corresponding acids or esters. In this case, acetic acid or aceticesters are advantageously hydrogenated to ethanol using hydrogen. Aceticacid may advantageously be obtained by carbonylation of methanol or byfermentation of carbohydrates.

Preferably, the ethanol feed treated in the process of the invention isan ethanol feed produced from renewable sources obtained from biomass.

Preheating Step a)

The ethanol feed optionally undergoes a step a) for preheating in a heatexchanger so as to produce a preheated ethanol feed by means of anexchange of heat with the effluent obtained from dehydration step e).When a pre-treatment step b) is carried out, the temperature at the endof step a) is in the range 70° C. to 130° C., preferably in the range110° C. to 130° C. The pressure of the ethanol feed is adjusted so thatit remains liquid at the end of preheating step a), at a value in therange 0.1 to 3 MPa.

Pre-Treatment Step b)

The ethanol feed, optionally preheated, optionally undergoes a step b)for pre-treatment so as to produce a pre-treated ethanol feed. Saidpre-treatment step can be used to eliminate the nitrogen-containingcompounds present in said preheated feed so as to limit deactivation ofthe dehydration catalyst placed downstream.

Said pre-treatment step b) is carried out over an acidic solid,preferably an acidic resin, and at a temperature in the range 70° C. to130° C., preferably in the range 110° C. to 130° C.

Said pre-treatment step b) can be used to eliminate impurities, basicand/or organic, and cationic species in order to obtain a pre-treatedethanol feed the level of impurities of which is compatible with thedehydration catalyst.

Pre-treatment over the acidic solid under the operating conditions ofthe invention means that between 3% by weight and 20% by weight,preferably between 8% and 12% by weight of the ethanol present in saidfeed can be converted into DEE, the percentage by weight beingdetermined with respect to the total weight of ethanol present in saidfeed at the inlet to pre-treatment step b).

The acidic solid includes all of the acidic solids known to the skilledperson: silica-aluminas, acid clays, zeolites, sulphated zirconias,acidic resins, etc. The essential fact is that the acidic solid has ahigh exchange capacity in order to capture as many of the basic andcationic species as possible, and a sufficiently high acid strength tocarry out the partial transformation of ethanol to DEE.

Acidic solids which are readily commercially available are clays treatedwith acids to make them acidic (such as montmorillonite) and zeoliteswith a silica to alumina ratio in the crystalline lattice of 2.5 to 100molar.

The acidic resin includes sulphonic groups grafted onto an organicsupport composed of aromatic and/or haloaliphatic chains. Preferably,the acidic solids have an exchange capacity of at least 0.1 mmol H⁺equivalent per gram.

The acidic resin is prepared by polymerization or copolymerization ofaromatic vinyl groups followed by a sulphonation, said aromatic vinylgroups being selected from styrene, vinyl toluene, vinyl naphthalene,vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinylxylene, said resin having a degree of cross-linking in the range 20% to35%, preferably in the range 25% to 35% and preferably equal to 30%, andan acid strength, measured by potentiometry during neutralization with aKOH solution, of 0.2 to 10 mmol H⁺ equivalent per gram, and preferably0.2 to 2.5 mmol H⁺ equivalent per gram.

Said acidic ion exchange resin contains between 1 and 2 terminalsulphonic groups per aromatic group. It is in the range 0.15 to 1.5 mmin size. The term “resin size” means the diameter of the smallest sphereencompassing the particle of resin. Resin size classes are measured bysieving on suitable sieves using a technique which is known to theskilled person.

A preferred resin is a resin constituted by copolymers of monovinylaromatics and polyvinyl aromatics, highly preferably a copolymer ofdivinylbenzene and polystyrene with a degree of cross-linking in therange 20% to 45%, preferably in the range 30% to 40%, and morepreferably equal to 35% and with an acid strength, representing thenumber of active sites of said resin, measured by potentiometry duringneutralization with a KOH solution, in the range 1 to 10 mmol H⁺equivalent per gram, and preferably in the range 3.5 to 6 mmol H⁺equivalent per gram. As an example, the resin is a TA801 resin sold byAxens.

The acidic solids may be regenerated from time to time once the exchangecapacity has been almost completely saturated by adsorption of basic andcationic species in situ or ex situ. In the case of inorganic acidicsolids such as clays and zeolites, regeneration may consist of simplyheating at high temperature in order to desorb the basic species in thepresence of an inert flow or a flow containing oxygen. The cations maybe removed by ion exchange. The acidic resins may be regenerated by ionexchange, typically by a treatment with an acid in the liquid phase. Theacidic solids may also be used once until saturation and replaced withfresh solid.

The acidic solid may be used alone or as a mixture with other types ofacidic solids. Mixtures of different acidic solids or sequences ofacidic solids may be used in order to optimize the adsorption capacityof the basic and cationic species and the capacity for partiallytransforming ethanol to DEE.

The pre-treatment described above may advantageously be supplemented bya pre-treatment using an anion exchange resin. This resin may, forexample, be a resin charged with sodium, or trimethylammoniumcharacterized by an exchange capacity measured in mg (OH⁻/litre). Thisresin may be Amberlite IRN78 resin, for example. This supplemental resincan be used to retain sulphate ions, SO₄ ²⁻, in order to extend the lifeof the catalyst.

Vaporization Step c)

The mixture comprising said pre-treated ethanol feed and at least aportion of the flow of diluting water comprising ethanol recycled inaccordance with recycling and vaporization step h) is known as the“vaporization feed”.

In accordance with the invention, the dehydration process comprises astep c) for vaporization of said vaporization feed so as to produce avaporized feed. Said vaporization is carried out by means of an exchangeof heat with the effluent obtained from dehydration step e) in a heatexchanger.

Preferably, said vaporization feed is introduced into said vaporizationstep c) at a pressure which is lower than the pressure of the effluentobtained from dehydration step e).

The pressure of said vaporization feed upstream of vaporization step c)is advantageously selected such that the temperature difference in theheat exchanger between the effluent obtained from the dehydration stepe) which is condensed and said vaporization feed which is evaporated isno less than 2° C., preferably no less than 3° C.

This difference in temperature in the heat exchanger is known as theapproach to thermal equilibrium.

Adjustment of said pressure upstream of the vaporization step c) is anessential criterion of the present invention. This pressure is selectedso as to be as high as possible, so that the difference between thecondensation temperature of the effluent from dehydration step e) andthe vaporization temperature of said vaporization feed is 2° C. or more,preferably 3° C. or more, in order to maximize the exchange of heatbetween said vaporization feed and said effluent obtained from thedehydration step e).

In a preferred arrangement in which the pre-treatment step b) is carriedout, the vaporization temperature of the vaporization feed at a givenpressure is reduced compared with that of a feed obtained by aconcatenation which does not include the pre-treatment step b). For agiven condensation temperature of the effluent from the dehydration stepe) and a fixed approach to thermal equilibrium, then, the pressureupstream of the vaporization step c) may be adjusted to a higher valuethan that which would have existed in a concatenation not including thepre-treatment step b).

Adjustment of said pressure upstream of the vaporization step c) to thehighest value possible, within the limits determined in the precedingparagraph, means that the energy necessary for compression during thecompression step d) of the process of the invention can be minimized.

Said vaporization feed is introduced into said vaporization step c) at apressure in the range 0.1 to 1.4 MPa, preferably in the range 0.2 to 0.6MPa.

Introducing said vaporization feed into the vaporization step c) at thisspecific pressure in the range 0.1 to 1.4 MPa, preferably in the range0.2 to 0.6 MPa, below the pressure of the effluent leaving the finalreactor of the dehydration step e), means that a vaporizationtemperature for said vaporization feed below the condensationtemperature of the effluent obtained from the last adiabatic reactor canbe exploited. Thus, the major portion of the latent heat of the aqueousphase of the effluent obtained from the last adiabatic reactor isrecovered in order to vaporize said vaporization feed without addingexternal heat. The whole of the enthalpy of vaporization of saidvaporization feed is thus exchanged with the enthalpy of condensation ofsaid effluent.

Compression and Superheating Step d)

In accordance with the invention, said vaporized feed undergoes acompression in a compression step d) so as to produce a compressed feed.Said compression step d) is advantageously carried out in any type ofcompressor known to the skilled person. In particular, compression stepd) is advantageously carried out in a compressor of the integrallygeared centrifugal compressor type or in a compressor comprising one ormore blowers with a radial wheel in series without intermediate cooling,or in a volumetric type compressor with or without lubrification.

In an arrangement in which step b) is carried out, said step b) cansurprisingly be carried out at a higher pressure upstream of step c);the degree of compression in step d) is reduced in order to obtain agiven pressure at the outlet from said step d), thereby reducing theenergy consumption of said step d).

The compression step d) can be used to produce a heat pump which isintegrated into said process, using the flows obtained from the processand not involving an external heat transfer fluid.

The combination of the specific operating conditions of step c) and stepd) means that the addition of external heat transfer fluid to theprocess to ensure vaporization of said vaporization feed can be avoidedby recovering the major portion of the latent heat of the aqueous phaseof the effluent obtained from the last adiabatic reactor in order tovaporize the vaporization feed. Thus, only flows obtained from theprocess are used.

The pressure of said compressed feed obtained from compression step d)is advantageously in the range 0.3 to 1.8 MPa, preferably in the range0.5 to 1.3 MPa. The pressure at the outlet for said feed is sufficientlyhigh for the condensation temperature of the effluent obtained from thelast reactor to be higher than the vaporization temperature for the feedentering step c), which is a necessary condition for step c) to befeasible.

Said compressed feed obtained from the compression step d) is optionallyheated in a single phase type gas exchanger by means of an exchange ofheat with the effluent obtained from the last adiabatic reactor of stepe). In said single phase type gas exchanger, said compressed feed issuperheated and the effluent obtained, in the gaseous state, from thelast adiabatic reactor of step e) is “desuperheated” without beingcondensed.

Said compressed feed is advantageously superheated to an outlettemperature in the range 250° C. to 420° C., preferably in the range280° C. to 410° C. At the outlet from said single phase type gasexchanger of the last adiabatic reactor of step e), the effluentobtained, in the gaseous state, is advantageously at a temperature inthe range 180° C. to 260° C.

Thus, using the various exchangers of the single phase gas andgas/liquid evaporator type, and vaporization at a pressure below thepressure of the effluent leaving the last reactor of said vaporizationfeed means that condensation of at least 60% of the water vapour presentin the effluent obtained from the last reactor of the dehydration stepe) can be carried out.

The single phase type gas exchanger is an exchanger the technology ofwhich is known to the skilled person, which can be used to minimizepressure drops while providing a large exchange surface. This lowpressure gas/gas exchange induces a low heat flux density through thewall of the exchanger (low transfer coefficient), which necessitateshaving a large exchange surface. In addition, the loss of pressure hasto be minimized in order to limit the load on the compressor of step d).As an example, this exchanger may be a pressurized plate exchanger in ashell, of the Packinox type provided by Alpha Laval.

Dehydration Step e)

In accordance with the invention, said compressed feed, optionallyheated, is mixed with the vaporized flow of diluting water comprisingethanol obtained from step h) for recycling and vaporization. Themixture is advantageously introduced into a furnace so as to bring it toan inlet temperature in at least one adiabatic reactor which iscompatible with the dehydration reaction temperature. It then undergoesa dehydration step e) in at least one adiabatic reactor containing atleast one fixed bed of dehydration catalyst and in which the dehydrationreaction takes place.

The mixture at the inlet to step e) for dehydration of the flowconstituted by the ethanol feed and the flow of diluting watercomprising ethanol obtained from step g) is produced in a manner suchthat at the end of mixing, the weight ratio of water to ethanol, knownas the dilution ratio, is in the range 1 to 4. Dilution is intended toreduce the partial pressures of ethanol in the reactor or reactors andto render the process more selective for ethylene. This weight ratio isadjusted by modifying the flow rate of treated water at the end of stepg) and/or by modifying the flow rate of the ethanol feed.

The dehydration step e) is advantageously carried out in one or tworeactors.

In the case in which step e) is carried out in an adiabatic reactor,said compressed feed, optionally heated, is advantageously introducedinto said reactor at an inlet temperature in the range 350° C. to 550°C., preferably in the range 400° C. to 500° C., and at an inlet pressurein the range 0.3 to 1.8 MPa, preferably in the range 0.4 to 0.8 MPa.

The effluent obtained from said adiabatic reactor of step e) isadvantageously at a temperature in the range 270° C. to 450° C.,preferably in the range 340° C. to 430° C., and with an outlet pressurein the range 0.2 to 1.6 MPa, preferably in the range 0.3 to 0.8 MPa.

In the case in which step e) is carried out in two adiabatic reactors,said compressed feed, also optionally heated, is advantageouslyintroduced into the first reactor at an inlet temperature in the range350° C. to 550° C., preferably in the range 370° C. to 500° C., and atan inlet pressure in the range 0.3 to 1.8 MPa, preferably in the range0.4 to 1.1 MPa.

The effluent obtained from the first adiabatic reactor advantageouslyleaves said first reactor at a temperature in the range 270° C. to 450°C., preferably in the range 290° C. to 390° C., and at a pressure in therange 0.3 to 1.7 MPa, preferably in the range 0.3 to 1.0 MPa.

Said effluent is advantageously introduced into a furnace in a mannersuch that the inlet temperature for said effluent into the secondadiabatic reactor is in the range 350° C. to 550° C., preferably in therange 400° C. to 500° C. Said effluent has an inlet pressure into saidsecond reactor which is advantageously in the range 0.3 to 1.7 MPa,preferably in the range 0.3 to 0.9 MPa.

The effluent obtained from the second adiabatic reactor leaves saidsecond adiabatic reactor at a temperature which is advantageously in therange 270° C. to 450° C., preferably in the range 340° C. to 430° C. Theoutlet pressure for said effluent obtained from the second adiabaticreactor is advantageously in the range 0.2 to 1.6 MPa, preferably in therange 0.3 to 0.8 MPa.

The inlet temperature for the reactor or reactors may advantageously begradually increased in order to prevent deactivation of the dehydrationcatalyst.

The dehydration reaction which takes place in at least one adiabaticreactor of step e) of the process of the invention advantageouslyoperates at a weight hourly space velocity in the range 0.1 to 20 h⁻¹,preferably in the range 0.5 to 15 h⁻¹. The weight hourly space velocityis defined as the ratio of the mass flow rate of the pure ethanol feedover the weight of catalyst.

The dehydration catalyst used in step e) is a catalyst which is known tothe skilled person. Said catalyst is preferably an amorphous acidcatalyst or a zeolitic acid catalyst.

In the case in which the dehydration catalyst used in step e) is azeolitic catalyst, said catalyst comprises at least one zeolite selectedfrom zeolites having at least pore openings containing 8, 10 or 12oxygen atoms (8 MR, 10 MR or 12 MR). It is in fact known to define thepore size of zeolites by the number of oxygen atoms forming the annularsection of the channels of the zeolites, known as the “member ring” orMR. Preferably, said zeolitic dehydration catalyst comprises at leastone zeolite with a structure type selected from structure types MFI,FAU, MOR, FER, SAPO, TON, CHA, EUO and BEA. Preferably, said zeoliticdehydration catalyst comprises a zeolite with structure type MFI, andmore preferably a ZSM-5 zeolite.

The zeolite employed in the dehydration catalyst used in step e) of theprocess of the invention may advantageously be modified by dealuminationor desilication using any dealumination or desilication method known tothe skilled person.

The zeolite employed in the dehydration catalyst used in step e) of theprocess of the invention or the final catalyst may advantageously bemodified by an agent of a nature such as to attenuate its total acidityand to improve its hydrothermal resistance properties. Preferably, saidzeolite or said catalyst advantageously comprises phosphorus, preferablyadded in the form of H₃PO₄, followed by steam treatment afterneutralization of the excess acid with a basic precursor such ascalcium, Ca, for example. Preferably, said zeolite has a phosphoruscontent in the range 1% to 4.5% by weight, preferably in the range 1.5%to 3.1% by weight with respect to the total catalyst weight.

Preferably, the dehydration catalyst used in step e) of the process ofthe invention is the catalyst described in patent applicationsWO/2009/098262, WO/2009/098267, WO/2009/098268, or WO/2009/098269.

In the case in which the dehydration catalyst used in step e) is anamorphous acid catalyst, said catalyst comprises at least one porousrefractory oxide selected from alumina, alumina activated by a depositof mineral acid, and silica-alumina.

Said amorphous or zeolitic dehydration catalyst used in step e) of theprocess of the invention may advantageously also comprise at least oneoxide type matrix also known as a binder. The term “matrix of theinvention” means an amorphous or crystalline matrix or a matrixcomprising amorphous and crystalline portions. Said matrix isadvantageously selected from elements of the group formed by clays (suchas natural clays, for example kaolin or bentonite), magnesia, aluminas,silicas, silica-aluminas, aluminates, titanium oxide, boron oxide,zirconia, aluminium phosphates, titanium phosphates, zirconiumphosphates and charcoal, used alone or as a mixture. Preferably, saidmatrix is selected from elements from the group formed by aluminas,silicas and clays.

Said dehydration catalyst used in step e) of the process of theinvention is advantageously shaped into the form of grains with variousshapes and dimensions. It is advantageously used in the form ofcylindrical or polylobed extrudates such as bilobes, trilobes, polylobesin a straight or twisted form, but may optionally be fabricated andemployed in the form of a crushed powder, tablets, rings, beads, wheelsor spheres. Preferably, said catalyst is in the form of extrudates.

Said dehydration catalyst used in step e) of the process of theinvention is advantageously deployed in at least one reactor, in a fixedbed or moving bed.

In step e) of the process of the invention, the catalysts used and theoperating conditions are selected so as to maximize the production ofethylene. The overall dehydration reactions involved in step e) of theprocess of the invention are as follows:

2C₂H₅OH→2CH₂═CH₂+2H₂O

CH₃CH₂OCH₂CH₃→2CH₂═CH₂+H₂O

Conversion of the ethanol feed in step e) is more than 90%, preferably95% and more preferably more than 99%.

A conversion of less than 90% has the effect of dropping the overallyield of the process, as a large quantity of DEE which is not convertedinto ethylene is lost in the downstream separation steps.

Conversion of the ethanol feed is defined, as a percentage, by thefollowing formula:

[1−(hourly mass of ethanol at outlet/hourly mass of ethanol atinlet)]×100

The hourly mass of ethanol at the inlet and outlet is measured in aconventional manner, for example by chromatography.

Step e) in which the dehydration reaction takes place is advantageouslycarried out in one or two reactors. A preferred reactor is a radialreactor functioning in upflow mode or downflow mode. During step e) ofthe process of the invention, transformation of the feed is accompaniedby deactivation of the dehydration catalyst by coking and/or byadsorption of inhibiting compounds. Thus, the dehydration catalystperiodically has to undergo a regeneration step. Preferably, the reactoris used in an alternating regeneration mode reactor, also known as aswing reactor, in order to alternate reaction phases with phases forregeneration of said dehydration catalyst. The aim of this regenerationtreatment is to burn off the organic deposits along with the speciescontaining nitrogen and sulphur contained at the surface and within saiddehydration catalyst. The pre-treatment step b) carried out in thisinvention can be used to reduce the quantity of impurities, basic andorganic, as well as the cationic species which would alter the cycletime of the catalyst. Eliminating these species thus means that thenumber of catalyst regenerations can be limited.

Regeneration of the dehydration catalyst used in said step e) isadvantageously carried out by oxidation of coke and inhibiting compoundsin a flow of air or in an air/nitrogen mixture, for example employingrecirculation of the combustion air with or without water in order todilute the oxygen and control the exothermic nature of the regeneration.In this case, the oxygen content at the inlet to the reactor canadvantageously be adjusted by adding air. Regeneration takes place at apressure in the range between atmospheric pressure and the reactionpressure.

The regeneration temperature is advantageously selected to be between400° C. and 600° C.; it may advantageously vary during regeneration. Theend of regeneration is detected when there is no more oxygenconsumption—an indication that combustion of the coke is complete.

The effluent obtained from the last adiabatic reactor of step e) isoptionally sent to a single phase type gas exchanger in which it is“desuperheated” without being condensed by exchange of heat with thecompressed feed obtained from step d), which is itself superheated.

Said “desuperheated” effluent is then advantageously sent to a secondgas/liquid type exchanger in which it is partially condensed by heatexchange in order to vaporize the vaporization feed.

Said effluent is then cooled further by heat exchange with the ethanolfeed during step a) for preheating the ethanol feed.

Separation Step f)

In accordance with the invention, the effluent obtained from the lastadiabatic reactor of step e) undergoes a step f) for separation into aneffluent comprising ethylene at a pressure of less than 1.6 MPa,preferably less than 0.8 MPa, and an effluent comprising water.

Step f) for separation of said effluent obtained from the last adiabaticreactor of step e) may advantageously be carried out using any methodknown to the skilled person such as, for example, a gas/liquidseparation zone, preferably a gas/liquid separation column.

The effluent comprising ethylene at a pressure of less than 1.6 MPa thenadvantageously undergoes a compression. Said compression allows thepressure of said effluent to be raised to a pressure which isadvantageously in the range 2 to 4 MPa, which is necessary for itspurification.

Purification Step g)

In accordance with the invention, the effluent comprising water obtainedfrom the separation step f) undergoes a purification step g). Thepurification step g) may be carried out using any purification methodwhich is known to the skilled person. By way of example, thepurification step g) may advantageously be carried out using ionexchange resins, by adding chemical agents such as sodium hydroxide oramines, for example, to adjust the pH, and by adding chemical agentssuch as, for example, polymerization inhibitors selected frombisulphites and surfactants to stabilize the products.

At least one flow of treated water and at least one flow of dilutingwater comprising ethanol are then separated. This ethanol corresponds toethanol which has not been converted in step e). The separation may beused to obtain a flow of ethanol-free treated water (less than 10% byweight of ethanol, preferably less than 1%), which limits the loss ofethanol, and may be carried out using any separation method known to theskilled person. By way of example, the separation may advantageously becarried out by distillation, using molecular sieves, steam or heatstripping or by solvent absorption using glycolated solvents, forexample.

A flow containing the light gases and ethanol, preferably acetaldehydeand methanol, may advantageously also be separated and recycled to stepf).

Recycling and Vaporization Step h)

According to the invention, the flow of diluting water comprising theethanol obtained from purification step g) undergoes a step h) forrecycling and vaporization.

At least a portion of the flow of diluting water comprising ethanolobtained from step g) undergoes vaporization using heat exchange with aquench flow obtained from the oxidation step j) in a vaporizationexchanger. The term “at least a portion” means a quantity which iscarefully selected so that at least 50%, preferably 80% is vaporized atthe outlet from the vaporization exchanger. The quench flow obtainedfrom step j) is partially, preferably completely, condensed during saidheat exchange then recycled to the oxidation step j).

The pressure of the step h) for recycling and vaporization of dilutingwater comprising ethanol is adjusted so as to allow injection withoutcompression of the vaporized flow of diluting water comprising ethanolupstream of step e). In the case in which the flow of diluting watercomprising ethanol is partially vaporized by heat exchange with thequench flow obtained from step j), the non-vaporized portion isseparated, decompressed, and recycled to the inlet to step c).

In a preferred arrangement, the flow of diluting water comprisingethanol obtained from step h) is injected into the inlet to step e).

In another preferred arrangement, the flow of diluting water comprisingethanol obtained from step h) is injected into the inlet to a singlephase gas/gas exchanger in step d) for compression and superheating.

The non-vaporized portion of the flow of diluting water comprisingethanol obtained from the purification step g) comprising thenon-vaporized fraction following heat exchange, as well as the portionwhich has not undergone heat exchange, is advantageously recycled andmixed with the pre-treated ethanol feed at the inlet to the vaporizationstep c).

By vaporizing a portion of the flow of diluting water comprising ethanolobtained from step g) and by recycling this vaporized portion upstreamof step e), dilution of ethanol at the inlet to step c) is minimized.Minimizing the dilution in this manner has the effect of reducing thetemperature and enthalpy of vaporization of the mixture in step c)compared with the situation in which all of the flow of diluting watercomprising ethanol is recycled to the inlet of step c), meaning thatevaporation of said mixture is facilitated and the energy requirementsin step c) are minimized.

The flows of recycled water act as a thermal reaction diluent.

Careful adjustment of the operating conditions for the ethyleneoxidation step j) and the dehydration step e) means that, by carryingout step h), the energy necessary for compression in step d) can belimited by avoiding compressing, in the vapour phase, the whole of theflow of diluting water comprising ethanol obtained from step g) bymaximizing the thermal integration of the process, while reducing thetotal consumption of water.

Compression Step i)

In accordance with the invention, the flow comprising ethylene obtainedfrom step f) then undergoes a compression using means known to theskilled person, for example a centrifugal or volumetric compressor,optionally multi-stage with intermediate cooling. The condensates mayadvantageously be mixed with the supply for the purification step g).This compression means that the effluent can be brought to a sufficientpressure to allow any purification steps which follow to take place, theoperating pressure of which is directly dependent on this compressor,ignoring the pressure drops. Preferably, the discharge pressure issufficiently high to allow direct injection of ethylene into theoxidation step j). This discharge pressure is in the range 1.1 to 5.1MPa, preferably in the range 1.6 to 3.6 MPa.

The compressed flow then advantageously undergoes purification usingmeans known to the skilled person, for example by one or moredistillations, optionally cryogenic, and/or using resin type capturemasses. Purification of the flow comprising ethylene depends on itsdestination, for example on the technology selected for oxidation stepi), or on the possibility or otherwise of producing ethylene withoutoxidizing it (at the same time as the production of ethylene oxide).

Oxidation Step j)

In accordance with the invention, at least a portion of the flowcomprising ethylene obtained from separation step f), compressed incompression step i) and optionally purified, is directed to an oxidationstep j). This oxidation step comprises a set of equipment, catalysts andchemical products which are known to the skilled person to allow theoxidation of ethylene to form ethylene oxide.

The oxidation catalyst is advantageously based on silver. The reactionselectivity may advantageously be optimized by controlling thetemperature of the reaction, and optionally by using methane dilution.This methane may derive either from storage or from a methane productionunit.

An adjustment of these parameters means that the quantity of heatreleased by the reaction can be controlled, and as a consequence theheat extracted via evaporation of the quench fluid and the heatrecoverable by condensation of this fluid in step h) can be controlled.

In accordance with the invention, the temperature of the ethyleneoxidation reaction is controlled at between 100° C. and 500° C.,preferably in the range 150° C. to 300° C., by means of vaporization ofa heat transfer flow, termed the quench flow, in the shell of one ormore multitube reactor(s), the reaction taking place in the tubes. Thequench flow preferably comprises water and may advantageously be aprocess water flow, distilled water flow or any other flow comprisingwater which is not or is only slightly fouling. The temperature of thereaction may, for example, be controlled by varying the pressure on theshell side of the reactor or reactors, which has the effect of varyingthe vaporization temperature of the quench fluid.

The vaporization pressure of the quench flow, and thus its vaporizationtemperature, is selected so as to maximize the density of the flow ofheat in the exchanger or exchangers of step h) for recycling andvaporization of diluting water comprising ethanol. The approach tothermal equilibrium of an exchange of heat is defined as the minimumabsolute value for the temperature difference between the two sides ofthe exchange of heat under consideration. The vaporization pressure ofthe quench flow is selected in a manner such that the approach tothermal equilibrium in the exchangers of step h) is a maximum,preferably more than 10° C., preferably more than 30° C. This means thatthe exchange surface and the cost of the vaporization exchanger(s) ofstep h) can be minimized.

The oxidation step j) produces at least one effluent comprising ethyleneoxide and at least one quench flow which is partially or completelyvaporized which is then directed to the step h) for recycling andvaporization.

DESCRIPTION OF THE FIGURES

FIG. 1 is a diagrammatic representation of the process for thedehydration of ethanol followed by oxidation of the ethylene produced inthe case of a concentrated feed of ethanol with a recycle of a flow ofdiluting water comprising ethanol obtained from step h) of the process.

The ethanol feed 1 is preheated in an exchanger a) with the effluentfrom the dehydration step e) which penetrates into the exchanger via theconduit 7. The preheated ethanol feed is then introduced into apre-treatment zone b) via the conduit 2. The pre-treated ethanol feed isthen mixed in the conduit 3 with the non-vaporized diluting watercomprising ethanol obtained from the recycling and vaporization zone h)which is recycled via the conduits 23 so as to act as a reactiondiluent. This mixture, constituting the vaporization feed, is introducedvia line 3 into the vaporization step c) in which said mixture undergoesheat exchange with the effluent obtained from the dehydration step e)which penetrates into the exchanger via the conduit 23 in a manner suchas to produce a vaporized feed. The latent heat, also known as theenthalpy of condensation, of the effluent obtained from the dehydrationstep e) is used to vaporize the vaporization feed without addingexternal heat.

The vaporized feed is then sent to the compression and superheating stepd) via the conduit 4. In the compression and superheating step d), saidvaporized feed is compressed and superheated and the effluent obtainedfrom the dehydration section e), in the gaseous state, is“desuperheated” without being condensed.

Said vaporized, compressed and superheated feed is then mixed with thevaporized flow of diluting water comprising ethanol obtained fromrecycling and vaporization step h) then introduced into step e) in whichit passes through a furnace or other equipment known to the skilledperson so as to bring it to a temperature which is compatible with thedehydration reaction temperature.

The effluent obtained from the dehydration section e) then undergoesthree successive exchanges described above in steps d), c) and a).

The effluent obtained from step a) is sent via the conduit 8 to theseparation section f) where it is separated into an effluent 9comprising ethylene and an effluent 14 comprising water.

The effluent comprising ethylene is sent to the compression step i) viathe line 9. Step i) can be used to bring said effluent comprisingethylene to a pressure sufficient for it to be purified, and preferablya pressure which is compatible with the supply pressure for theoxidation section j). Step i) can also be used to purify the effluentcomprising ethylene by separating at least one flow 17 comprising thelight gases and a flow 18 of water and unreacted ethanol recycled to thepurification step g).

The compressed and purified effluent 10 comprising ethylene obtainedfrom step i) is sent to the oxidation section j) where it undergoes aconversion into ethylene oxide. Step j) can be used to form a flow 11comprising ethylene oxide. The heat of reaction of step j) can be usedto vaporize the quench fluid 13 which is then recycled via the conduit12 to the recycling and vaporization step h).

The effluent comprising water obtained from step f) is sent to apurification step g) via the conduit 14. At least one flow of dilutingwater comprising unconverted ethanol and at least one flow 19 of treatedwater are separated. A flow 21 containing the light gases andunconverted ethanol is also separated and recycled to step f).

The whole of said flow of diluting water comprising unconverted ethanolobtained from the purification step g) is sent to the recycling andvaporization step h), where it is divided into two flows 15 and 23. Theflow 15 is optionally pumped at a pressure which is compatible withinjection of the flow 16 upstream of step e), then is partially orcompletely vaporized by means of an exchange of heat using thecondensation of the vaporized quench fluid 12 obtained from theoxidation step j). The condensed quench fluid is recycled to theoxidation section j) via the conduit 13. The vaporized effluent 16 fromstep h) is incorporated with the pre-treated ethanol feed upstream ofthe dehydration step e).

The non-vaporized portion 23 obtained from step h) is recycled upstreamof the vaporization step c). In the case in which vaporization of theflow 15 is not complete, the liquid portion may advantageously bedecompressed and mixed with flow 23.

The following examples illustrate the invention without limiting itsscope.

EXAMPLES Example 1 In Accordance with the Invention

Example 1 illustrates a process in accordance with the invention.

The ethanol feed under consideration was produced by fermentation ofwheat without extracting the gluten, employing a dry milling process.

Step a)

Said ethanol feed was introduced into an exchanger E1 at a flow rate of45 664 kg/h at a pressure of 1.15 MPa and was heated, remaining in theliquid phase, to a temperature of 120° C. against the effluent obtainedfrom the adiabatic reactor of step e).

Step b)

The heated ethanol feed was pre-treated over a TA801 resin in order toeliminate traces of nitrogen-containing compounds. During thispre-treatment, a portion of the ethanol was converted into DEE. Thecharacteristics of the unrefined and pre-treated ethanol feed are givenin Table 1.

TABLE 1 Characteristics of ethanol feed before and after pre-treatment(percentages by weight) ETHANOL AFTER ETHANOL FEED PRETREATMENT ETHANOL91.2% 82.1% H2O  8.7% 10.5% DEE   0%  7.3% NITROGEN- 0.005%  0.000% CONTAINING COMPOUNDS

Step c)

The vaporization feed, constituted by the pre-treated ethanol feed mixedwith 85380 kg/h of diluting water and unconverted ethanol recycled inaccordance with step h), was introduced into an exchanger E2 at apressure of 0.37 MPa. The vaporization feed entered the exchanger E2 at113° C. and was thus already vaporized to an extent of 8% by weight. Thepressure at the inlet to the exchanger E2 was adjusted so that theapproach to thermal equilibrium with the flow obtained from the lastadiabatic reactor of step e) was a minimum of 15° C.

In step c), the majority of the latent heat of the aqueous phase of theeffluent obtained from the last adiabatic reactor of step e) wasrecovered in order to vaporize the vaporization feed without addingexternal heat. Thus, 51.9 MW was exchanged between said vaporizationfeed and said effluent.

Step d)

The vaporized feed was then compressed in an integrally gearedcentrifugal compressor K1 in a manner such that the pressure of saidvaporized feed was equal to 0.695 MPa at the end of compression.

The compressed feed was then heated in a single phase type gas exchangerE3 using an exchange of heat with the effluent obtained from theadiabatic reactor of step e). In said single phase type gas exchanger,said compressed feed was superheated to a temperature of 419° C. and theeffluent obtained, in the gaseous state, from the last adiabatic reactorof step e) was “desuperheated” without being condensed and had atemperature of 273° C.

Step e)

Dehydration step e) comprised two furnaces and two adiabatic reactors inseries.

Said feed, compressed and heated in said single phase type gasexchanger, was then mixed with 55590 kg/h of vaporized diluting waterobtained from the recycling and vaporization step h). This mixture wasthen introduced into a furnace so as to bring it to an inlet temperaturefor the first adiabatic reactor of step e) which was compatible with thetemperature of the highly endothermic dehydration and conversionreaction of DEE to ethylene, i.e. to a temperature of 470° C. The outlettemperature from the last adiabatic reactor of step e) was 420° C.

The degree of dilution at the reactor inlet (i.e. the water/ethanolweight ratio) was 3.

Trapping the nitrogen-containing compounds in pre-treatment step b)meant that the inlet temperature to the first adiabatic reactor of stepe) could be considerably reduced.

Said compressed and heated feed was introduced into the first adiabaticreactor at an inlet pressure of 0.595 MPa. The pressure of the effluentat the outlet from the last adiabatic reactor from step e) was 0.500MPa. The dehydration step e) was operated at a weight hourly spacevelocity of 7 h⁻¹.

The adiabatic reactor contained a fixed bed of dehydration catalyst,said catalyst comprising 80% by weight of ZSM-5 zeolite treated withH₃PO₄ such that the phosphorus P content was 3% by weight.

The conversion of the ethanol feed in step e) was 95%.

Step f)

The effluent obtained from the last adiabatic reactor of step e) thenunderwent three exchanges of heat as described above and was sent to agas/liquid separation column. An effluent comprising ethylene at apressure of 0.36 MPa was then separated, along with an effluentcomprising water. This separation was carried out using a gas/liquidseparation column with a recycle of the water produced at the columnbottom to the head of the column and after cooling and injecting aneutralizing agent.

Step g)

A flow of diluting water comprising unconverted ethanol 15 as well as aflow containing light gases 21 was then separated by conventional lowpressure distillation of the impure water. A flow of treated water wasseparated 19 and constituted the purge for the process; its partialwater flow rate corresponded to the water formed by the dehydrationreaction of step e).

Step h)

The flow of diluting water comprising unconverted ethanol was dividedinto two flows, 23 and 15. The flow 15 was sent to a reboiler in orderto vaporize it using the heat supplied by the flow 12 obtained from theoxidation section j). 35 MW were exchanged in this step and could beused to vaporize 100% by weight of the flow 15. The hot flow obtainedfrom step j) was condensed and recycled to the reactor for the oxidationstep j). This flow was solely constituted by 65914 kg/h of distilledwater; its inlet pressure into step h) was 1.5 MPa at its dew pointtemperature, in this case 200° C.

The vaporized flow 16 was recycled upstream of the step e) and theliquid flow 23 was recycled upstream of step c).

Step i)

The effluent comprising ethylene then underwent compression to raise itspressure to 2.78 MPa before its final purification carried out withcryogenic distillation. A flow 17 of light gases was separated from thehead of this column and a flow of condensates comprising water andethanol was recycled to step g).

Step j)

The purified ethylene effluent obtained from step i) was sent to theoxidation section j). The operating conditions for this section were asfollows:

-   -   conversion per pass: 16%;    -   selectivity: 85% molar for ethylene;    -   O₂/ethylene ratio=7;    -   supply pressure to oxidation reactor: 20 barg;    -   reactor inlet temperature: 230° C.;    -   exothermicity of reaction: 40° C.

The exothermicity of the reaction, and thus its selectivity, was limitedbecause of a large dilution of the reaction medium by methane which wasrecycled to the process. The methane was purified using an absorptioncolumn employing a solution of potassium carbonate.

The reactor was a multitube reactor for generating 65914 kg/h ofsaturated vapour at a pressure of 1.5 MPa. This vapour was sent to therecycling and vaporization section h) by means of the flow 12.

The ethylene oxide was separated in a water absorption column followedby a stripping column.

A mixture of additional oxygen and methane was injected into theoxidation section j) via the conduit 22.

The purified ethylene oxide was extracted in the flow 11 before beingsent to an ethylene glycol conversion section.

The various flows, in kg/h, are consigned to Table 2:

TABLE 2 Composition of principal flows Current 1 2 3 4 5 6 Temperature °C. 25 120 113 132 419 420 Water kg/h 3993 3993 88133 88132.6 88132.6158602 Ethanol kg/h 41671 41671 42903 42903 42903 2187 DEE kg/h 8 8 8 14Ethylene kg/h 0 0 0 25228 kg/h Other (light + kg/h 603oxygen-containing) Total 45664 45664 131044 131044 131044 186634 Current8 19 16 12 13 17 Temperature ° C. 135 25 170 200 35 25 Water kg/h 15860219680 54782 65914 65914 Ethanol kg/h 2187 103 802 50 DEE kg/h 14 6Ethylene kg/h 25228 Ethylene oxide kg/h Other (light + kg/h 603 100 503oxygen-containing) Total 186634 19883 55590 65914 65914 553 Current 23 910 11 22 Temperature ° C. 143 35 25 35 35 Water kg/h 84140 300 Ethanolkg/h 1232 150 DEE kg/h 8 0 Ethylene kg/h 25228 25228 Ethylene oxide kg/h33438 Other (light + kg/h 503 8210 oxygen-containing) Total 85380 2618125228 33438 8210

The light compounds are C3 and C4 hydrocarbon compounds as well ashydrogen and carbon dioxide.

The selectivity of the process for the transformation of ethanol intoethylene was 99%.

It was calculated as follows: (ethylene contained in effluent comprisingethylene)/(0.61*quantity of converted ethanol), where the quantity ofconverted ethanol is the ethanol contained in the ethanol feed beforepre-treatment minus the ethanol contained in the purged flows of waterand in the effluent comprising ethylene. 0.61 g is the maximum quantityof ethylene obtained by dehydrating 1 g of pure ethanol.

This example shows that the exchanger E3 can be used to recover 35 MW ofthermal energy obtained from the oxidation section in order to vaporizediluting water from the dehydration section. The thermal energy of thisheat source was substantially higher (200° C. in this example) than thetemperature for the vaporization of the diluting water comprisingethanol (approximately 135° C. in this example), and so the density ofthe heat flux passing through the exchanger E3 was high compared with aprior art process where vaporization of the diluting water is carriedout by exchange with the effluent from the dehydration step e).

In addition, the fact of using a closed loop of distilled water torecover heat produced in the oxidation step j) has the advantage of notnecessitating a purification section for the quench water beforerecycling it (compared with a process in which the diluting water isused directly as a quench fluid in the shell of the oxidation reactorbefore being injected into the dehydration section).

The fact that the whole of the flow of diluting water comprising ethanolobtained from step g) is not compressed means that the power necessaryfor compression in step d) can be reduced by 40% compared with a priorart process in which the flow of diluting water comprising ethanolobtained from step g) is recycled in its entirety upstream of thevaporization step c).

1. A process for the dehydration of an ethanol feed to ethylene then oxidation of the ethylene to ethylene oxide, comprising: a) an optional step for preheating said ethanol feed to a temperature in the range 70° C. to 130° C. by heat exchange with the effluent obtained from step e); b) an optional step for pre-treatment of the ethanol feed over an acidic solid operating at a temperature in the range 70° C. to 130° C. in order to produce a pre-treated ethanol feed; c) a step for vaporization of a vaporization feed comprising said pre-treated ethanol feed and at least a portion of the flow of diluting water comprising ethanol recycled in accordance with step h) to an exchanger by means of an exchange of heat with the effluent obtained from the last reactor of step e), said vaporization feed being introduced into said vaporization step at a pressure in the range 0.1 to 1.4 MPa so as to produce a vaporized feed; d) a step for compressing and superheating said vaporized feed in a compressor so as to produce a feed which is compressed and at the temperature for supplying to the dehydration section e); e) a step for dehydration of a mixture comprising said compressed feed and the vaporized flow of diluting water comprising the ethanol obtained from step h), said mixture having a ratio of water to ethanol in the range 1 to 4 by weight, in at least one adiabatic reactor containing at least one dehydration catalyst and in which the dehydration reaction takes place, operating at an inlet temperature in the range 350° C. to 550° C. and at an inlet pressure in the range 0.3 to 1.8 MPa; f) a step for separating the effluent obtained from the last adiabatic reactor of step e) into an effluent comprising ethylene at a pressure of less than 1.6 MPa and an effluent comprising water; g) a step for purification of at least a portion of the effluent comprising water obtained from step f) and separation into at least a flow of treated water and a flow of diluting water comprising ethanol; h) a step for recycling and vaporization of at least a portion of the flow of diluting water comprising ethanol obtained from step g) by partial or total vaporization in an exchanger using heat exchange with a quench flow obtained from oxidation step j) in a vaporization exchanger, said quench flow, cooled, then being recycled to the reactor or reactors of oxidation step j), said flow of at least partially vaporized diluting water comprising ethanol being recycled upstream of step e), the non-vaporized portion of the flow of treated water obtained from step g) being recycled to the inlet to vaporization step c); i) a step for compression of the effluent comprising ethylene obtained from step f); j) a step for oxidation of the ethylene comprised in the effluent comprising ethylene obtained from step i) into ethylene oxide, this oxidation step comprising at least one tubular oxidation reactor cooled by vaporization of said quench flow obtained from step h), said quench flow which has thus been reheated being recycled to step h).
 2. The process according to claim 1, in which the flow of diluting water comprising ethanol obtained from step h) is injected into the inlet to step e).
 3. The process according to claim 1, in which the flow of diluting water comprising ethanol obtained from step h) is injected into the inlet to a single phase gas/gas exchanger in compression and superheating step d).
 4. The process according to claim 1, in which said compressed feed is heated in a single phase type gas exchanger by means of an exchange of heat with the effluent obtained from the last adiabatic reactor of step e).
 5. The process according to claim 1, in which the compressed flow obtained from said step i) undergoes a purification.
 6. The process according to claim 1, in which the portion of the flow of diluting water comprising ethanol obtained from step g) exchanging heat with the quench flow is such that at least 50% of said flow of diluting water is vaporized at the outlet from said vaporization exchanger.
 7. The process according to claim 1, in which the pressure of the compressed feed is in the range 0.3 to 1.8 MPa.
 8. The process according to claim 1, in which the effluent obtained from the last adiabatic reactor of step e) is at a temperature in the range 270° C. to 450° C., and at a pressure in the range 0.2 to 1.6 MPa.
 9. The process according to claim 1, in which the dehydration step e) is carried out in one or two reactors.
 10. The process according to claim 1, in which said dehydration catalyst used in step e) is an amorphous acid catalyst or a zeolitic acid catalyst.
 11. The process according to claim 1, in which said ethanol feed comprises a percentage by weight of ethanol in the range 35% to 99.9% by weight.
 12. The process according to claim 1, in which the optional step b) for pre-treatment is supplemented by a pre-treatment using an anion exchange resin. 